Process for recovering ethanol with vapor separation

ABSTRACT

Recovery of ethanol from a crude ethanol product obtained from the hydrogenation of acetic acid. Separation and purification processes of the crude ethanol products are employed to allow recovery of ethanol and remove impurities.

CROSS REFERENCE TO RELATED APPLICATIONS

This application is a continuation application of U.S. application Ser.No. 12/852,227, filed Aug. 6, 2010, which claims priority to U.S.Provisional Application No. 61/300,815, filed on Feb. 2, 2010, and U.S.Provisional Application No. 61/332,696, filed on May 7, 2010, the entirecontents and disclosures of which are incorporated herein by reference.

FIELD OF THE INVENTION

The present invention relates generally to processes for producingethanol and, in particular, to processes for producing ethanol from thehydrogenation of acetic acid.

BACKGROUND OF THE INVENTION

Ethanol for industrial use is conventionally produced from petrochemicalfeed stocks, such as oil, natural gas, or coal, from feed stockintermediates, such as syngas, or from starchy materials or cellulosematerials, such as corn or sugar cane. Conventional methods forproducing ethanol from petrochemical feed stocks, as well as fromcellulose materials, include the acid-catalyzed hydration of ethylene,methanol homologation, direct alcohol synthesis, and Fischer-Tropschsynthesis. Instability in petrochemical feed stock prices contributes tofluctuations in the cost of conventionally produced ethanol, making theneed for alternative sources of ethanol production all the greater whenfeed stock prices rise. Starchy materials, as well as cellulosematerial, are converted to ethanol by fermentation. However,fermentation is typically used for consumer production of ethanol forfuels or consumption. In addition, fermentation of starchy or cellulosematerials competes with food sources and places restraints on the amountof ethanol that can be produced for industrial use.

Ethanol production via the reduction of alkanoic acids and/or othercarbonyl group-containing compounds has been widely studied, and avariety of combinations of catalysts, supports, and operating conditionshave been mentioned in the literature. During the reduction of alkanoicacid, e.g., acetic acid, other compounds are formed with ethanol or areformed in side reactions. These impurities limit the production andrecovery of ethanol from such reaction mixtures. For example, duringhydrogenation, esters are produced that together with ethanol and/orwater form azeotropes, which are difficult to separate. In addition whenconversion is incomplete, unreacted acid remains in the crude ethanolproduct, which must be removed to recover ethanol.

Therefore, a need remains for improving the recovery of ethanol from acrude product obtained by reducing alkanoic acids, such as acetic acid,and/or other carbonyl group-containing compounds.

SUMMARY OF THE INVENTION

In a first embodiment, the present invention is directed to a processfor recovering ethanol, comprising hydrogenating an acetic acid feedstream with excess hydrogen in a reactor in the presence of a catalystto form a crude ethanol product; separating at least a portion of thecrude ethanol product in a first flasher into a first vapor stream andan intermediate stream; separating at least a portion of theintermediate stream in a second flasher into a second vapor stream and aliquid stream; and recovering ethanol from the liquid stream.

In a second embodiment, the present invention is directed to a processfor recovering ethanol, comprising providing a crude ethanol productcomprising ethanol, water, ethyl acetate, and acetaldehyde; separatingat least a portion of the crude ethanol product in a first flasher intoa first vapor stream and an intermediate stream; separating at least aportion of the intermediate stream in a second flasher into a secondvapor stream and a liquid stream; and recovering ethanol from the liquidstream.

In an another embodiment, the present invention is directed to a processfor recovering ethanol, comprising hydrogenating an acetic acid feedstream in the presence of a catalyst to form a crude ethanol product;separating at least a portion of the crude ethanol product in a firstcolumn into a first distillate comprising ethanol, water and ethylacetate, and a first residue comprising acetic acid; separating at leasta portion of the first distillate in a second column into a seconddistillate comprising ethyl acetate and a second residue comprisingethanol and water; and separating at least a portion of the secondresidue in a third column into a third distillate comprising ethanol anda third residue comprising water.

In yet another embodiment, the present invention is directed to aprocess for recovering ethanol, comprising providing a crude ethanolproduct comprising ethanol, water, acetic acid, and ethyl acetate;separating the crude ethanol product in a first column into a firstdistillate comprising ethanol, water and ethyl acetate, and a firstresidue comprising acetic acid; separating the first distillate in asecond column into a second distillate comprising ethyl acetate and asecond residue comprising ethanol and water; and separating the secondresidue in a third column into a third distillate comprising ethanol anda third residue comprising water.

BRIEF DESCRIPTION OF DRAWINGS

The invention is described in detail below with reference to theappended drawings, wherein like numerals designate similar parts.

FIG. 1A is a schematic diagram of a hydrogenation system in accordancewith one embodiment of the present invention.

FIG. 1B is a schematic diagram of the system shown in FIG. 1A with areturn of the distillate of a second column to the reaction zone inaccordance with one embodiment of the present invention.

FIG. 2 is a schematic diagram of the reaction zone in accordance withone embodiment of the present invention.

DETAILED DESCRIPTION OF THE INVENTION

The present invention relates to processes for recovering ethanolproduced by a hydrogenation process comprising hydrogenating acetic acidin the presence of a catalyst. In particular, the present inventionrelates to recovering ethanol from a crude ethanol product preferablyproduced by the hydrogenation process. Embodiments of the presentinvention beneficially may be used in applications for recovery ofethanol on an industrial scale.

Hydrogenation Process

The hydrogenation of acetic acid to form ethanol and water may berepresented by the following reaction:

Suitable hydrogenation catalysts include catalysts comprising a firstmetal and optionally one or more of a second metal, a third metal oradditional metals, optionally on a catalyst support. The first andoptional second and third metals may be selected from Group IB, IIB,IIIB, IVB, VB, VIIB, VIIB, VIII transitional metals, a lanthanide metal,an actinide metal or a metal selected from any of Groups IIIA, IVA, VA,and VIA. Preferred metal combinations for some exemplary catalystcompositions include platinum/tin, platinum/ruthenium, platinum/rhenium,palladium/ruthenium, palladium/rhenium, cobalt/palladium,cobalt/platinum, cobalt/chromium, cobalt/ruthenium, silver/palladium,copper/palladium, nickel/palladium, gold/palladium, ruthenium/rhenium,and ruthenium/iron. Exemplary catalysts are further described in U.S.Pat. No. 7,608,744 and U.S. Publication No. 2010/0029995, the entiretiesof which are incorporated herein by reference. Additional catalysts aredescribed in U.S. Pub. No. 2010/0197485, the entirety of which isincorporated herein by reference.

In one exemplary embodiment, the catalyst comprises a first metalselected from the group consisting of copper, iron, cobalt, nickel,ruthenium, rhodium, palladium, osmium, iridium, platinum, titanium,zinc, chromium, rhenium, molybdenum, and tungsten. Preferably, the firstmetal is selected from the group consisting of platinum, palladium,cobalt, nickel, and ruthenium. More preferably, the first metal isselected from platinum and palladium. When the first metal comprisesplatinum, it is preferred that the catalyst comprises platinum in anamount less than 5 wt. %, e.g., less than 3 wt. % or less than 1 wt. %,due to the high demand for platinum.

As indicated above, the catalyst optionally further comprises a secondmetal, which typically would function as a promoter. If present, thesecond metal preferably is selected from the group consisting of copper,molybdenum, tin, chromium, iron, cobalt, vanadium, tungsten, palladium,platinum, lanthanum, cerium, manganese, ruthenium, rhenium, gold, andnickel. More preferably, the second metal is selected from the groupconsisting of copper, tin, cobalt, rhenium, and nickel. More preferably,the second metal is selected from tin and rhenium.

If the catalyst includes two or more metals, e.g., a first metal and asecond metal, the first metal optionally is present in the catalyst inan amount from 0.1 to 10 wt. %, e.g., from 0.1 to 5 wt. %, or from 0.1to 3 wt. %. The second metal preferably is present in an amount from 0.1to 20 wt. %, e.g., from 0.1 to 10 wt. %, or from 0.1 to 5 wt. %. Forcatalysts comprising two or more metals, the two or more metals may bealloyed with one another or may comprise a non-alloyed metal solution ormixture.

The preferred metal ratios may vary depending on the metals used in thecatalyst. In some exemplary embodiments, the mole ratio of the firstmetal to the second metal is from 10:1 to 1:10, e.g., from 4:1 to 1:4,from 2:1 to 1:2, from 1.5:1 to 1:1.5 or from 1.1:1 to 1:1.1.

The catalyst may also comprise a third metal selected from any of themetals listed above in connection with the first or second metal, solong as the third metal is different from the first and second metals.In preferred aspects, the third metal is selected from the groupconsisting of cobalt, palladium, ruthenium, copper, zinc, platinum, tin,and rhenium. More preferably, the third metal is selected from cobalt,palladium, and ruthenium. When present, the total weight of the thirdmetal preferably is from 0.05 to 4 wt. %, e.g., from 0.1 to 3 wt. %, orfrom 0.1 to 2 wt. %.

In addition to one or more metals, the exemplary catalysts furthercomprise a support or a modified support, meaning a support thatincludes a support material and a support modifier, which adjusts theacidity of the support material. The total weight of the support ormodified support, based on the total weight of the catalyst, preferablyis from 75 wt. % to 99.9 wt. %, e.g., from 78 wt. % to 97 wt. %, or from80 wt. % to 95 wt. %. In preferred embodiments that use a modifiedsupport, the support modifier is present in an amount from 0.1 wt. % to50 wt. %, e.g., from 0.2 wt. % to 25 wt. %, from 0.5 wt. % to 15 wt. %,or from 1 wt. % to 8 wt. %, based on the total weight of the catalyst.

Suitable support materials may include, for example, stable metaloxide-based supports or ceramic-based supports. Preferred supportsinclude silicaceous supports, such as silica, silica/alumina, a GroupIIA silicate such as calcium metasilicate, pyrogenic silica, high puritysilica, and mixtures thereof. Other supports may include, but are notlimited to, iron oxide, alumina, titania, zirconia, magnesium oxide,carbon, graphite, high surface area graphitized carbon, activatedcarbons, and mixtures thereof.

In the production of ethanol, the catalyst support may be modified witha support modifier. Preferably, the support modifier is a basic modifierthat has a low volatility or no volatility. Such basic modifiers, forexample, may be selected from the group consisting of: (i) alkalineearth oxides, (ii) alkali metal oxides, (iii) alkaline earth metalmetasilicates, (iv) alkali metal metasilicates, (v) Group IIB metaloxides, (vi) Group IIB metal metasilicates, (vii) Group IIIB metaloxides, (viii) Group IIIB metal metasilicates, and mixtures thereof. Inaddition to oxides and metasilicates, other types of modifiers includingnitrates, nitrites, acetates, and lactates may be used. Preferably, thesupport modifier is selected from the group consisting of oxides andmetasilicates of any of sodium, potassium, magnesium, calcium, scandium,yttrium, and zinc, as well as mixtures of any of the foregoing.Preferably, the support modifier is a calcium silicate, and morepreferably calcium metasilicate (CaSiO₃). If the support modifiercomprises calcium metasilicate, it is preferred that at least a portionof the calcium metasilicate is in crystalline form.

A preferred silica support material is SS61138 High Surface Area (HSA)Silica Catalyst Carrier from Saint-Gobain NorPro. The Saint-GobainNorPro SS61138 silica contains approximately 95 wt. % high surface areasilica; a surface area of about 250 m²/g; a median pore diameter ofabout 12 nm; an average pore volume of about 1.0 cm³/g as measured bymercury intrusion porosimetry and a packing density of about 0.352 g/cm³(22 lb/ft³).

A preferred silica/alumina support material is KA-160 (Sud Chemie)silica spheres having a nominal diameter of about 5 mm, a density ofabout 0.562 g/ml, in absorptivity of about 0.583 g H₂O/g support, asurface area of about 160 to 175 m²/g, and a pore volume of about 0.68ml/g.

As will be appreciated by those of ordinary skill in the art, supportmaterials are selected such that the catalyst system is suitably active,selective and robust under the process conditions employed for theformation of ethanol.

The metals of the catalysts may be dispersed throughout the support,coated on the outer surface of the support (egg shell) or decorated onthe surface of the support.

The catalyst compositions suitable for use with the present inventionpreferably are formed through metal impregnation of the modifiedsupport, although other processes such as chemical vapor deposition mayalso be employed. Such impregnation techniques are described in U.S.Pat. No. 7,608,744, U.S. Publication No. 2010/0029995, and U.S.application Ser. No. 12/698,968, referred to above, the entireties ofwhich are incorporated herein by reference.

Some embodiments of the process of hydrogenating acetic acid to formethanol according to one embodiment of the invention may include avariety of configurations using a fixed bed reactor or a fluidized bedreactor, as one of skill in the art will readily appreciate. In manyembodiments of the present invention, an “adiabatic” reactor can beused; that is, there is little or no need for internal plumbing throughthe reaction zone to add or remove heat. In other embodiments, radialflow reactor or reactors may be employed, or a series of reactors may beemployed with or with out heat exchange, quenching, or introduction ofadditional feed material. Alternatively, a shell and tube reactorprovided with a heat transfer medium may be used. In many cases, thereaction zone may be housed in a single vessel or in a series of vesselswith heat exchangers therebetween.

In preferred embodiments, the catalyst is employed in a fixed bedreactor, e.g., in the shape of a pipe or tube, where the reactants,typically in the vapor form, are passed over or through the catalyst.Other reactors, such as fluid or ebullient bed reactors, can beemployed. In some instances, the hydrogenation catalysts may be used inconjunction with an inert material to regulate the pressure drop of thereactant stream through the catalyst bed and the contact time of thereactant compounds with the catalyst particles.

The hydrogenation reaction may be carried out in either the liquid phaseor vapor phase. Preferably, the reaction is carried out in the vaporphase under the following conditions. The reaction temperature may rangefrom 125° C. to 350° C., e.g., from 200° C. to 325° C., from 225° C. to300° C., or from 250° C. to 300° C. The pressure may range from 10 kPato 3000 kPa (about 1.5 to 435 psi), e.g., from 50 kPa to 2300 kPa, orfrom 100 kPa to 1500 kPa. The reactants may be fed to the reactor at agas hourly space velocity (GHSV) of greater than 500 hr⁻¹, e.g., greaterthan 1000 hr⁻¹, greater than 2500 hr⁻¹ or even greater than 5000 hr⁻¹.In terms of ranges the GHSV may range from 50 hr⁻¹ to 50,000 hr⁻¹, e.g.,from 500 hr⁻¹ to 30,000 hr⁻¹, from 1000 hr⁻¹ to 10,000 hr⁻¹, or from1000 hr⁻¹ to 6500 hr⁻¹.

The hydrogenation optionally is carried out at a pressure justsufficient to overcome the pressure drop across the catalytic bed at theGHSV selected, although there is no bar to the use of higher pressures,it being understood that considerable pressure drop through the reactorbed may be experienced at high space velocities, e.g., 5000 hr⁻¹ or6,500 hr⁻¹.

Although the reaction consumes two moles of hydrogen per mole of aceticacid to produce one mole of ethanol, the actual molar ratio of hydrogento acetic acid in the feed stream may vary from about 100:1 to 1:100,e.g., from 50:1 to 1:50, from 20:1 to 1:2, or from 12:1 to 1:1. Mostpreferably, the molar ratio of hydrogen to acetic acid is greater than2:1, e.g., greater than 4:1 or greater than 8:1.

Contact or residence time can also vary widely, depending upon suchvariables as amount of acetic acid, catalyst, reactor, temperature andpressure. Typical contact times range from a fraction of a second tomore than several hours when a catalyst system other than a fixed bed isused, with preferred contact times, at least for vapor phase reactions,of from 0.1 to 100 seconds, e.g., from 0.3 to 80 seconds or from 0.4 to30 seconds.

The raw materials, acetic acid and hydrogen, used in connection with theprocess of this invention may be derived from any suitable sourceincluding natural gas, petroleum, coal, biomass, and so forth. Asexamples, acetic acid may be produced via methanol carbonylation,acetaldehyde oxidation, ethylene oxidation, oxidative fermentation, andanaerobic fermentation. As petroleum and natural gas prices fluctuate,becoming either more or less expensive, methods for producing aceticacid and intermediates such as methanol and carbon monoxide fromalternate carbon sources have drawn increasing interest. In particular,when petroleum is relatively expensive compared to natural gas, it maybecome advantageous to produce acetic acid from synthesis gas (“syngas”) that is derived from any available carbon source. U.S. Pat. No.6,232,352, the disclosure of which is incorporated herein by reference,for example, teaches a method of retrofitting a methanol plant for themanufacture of acetic acid. By retrofitting a methanol plant, the largecapital costs associated with CO generation for a new acetic acid plantare significantly reduced or largely eliminated. All or part of the syngas is diverted from the methanol synthesis loop and supplied to aseparator unit to recover CO and hydrogen, which are then used toproduce acetic acid. In addition to acetic acid, such a process can alsobe used to make hydrogen which may be utilized in connection with thisinvention.

Methanol carbonylation processes suitable for production of acetic acidare described in U.S. Pat. Nos. 7,208,624, 7,115,772, 7,005,541,6,657,078, 6,627,770, 6,143,930, 5,599,976, 5,144,068, 5,026,908,5,001,259, and 4,994,608, the disclosure of which is incorporated hereinby reference. Optionally, the production of ethanol may be integratedwith such methanol carbonylation processes.

U.S. Pat. No. RE 35,377 also incorporated herein by reference, providesa method for the production of methanol by conversion of carbonaceousmaterials such as oil, coal, natural gas and biomass materials. Theprocess includes hydrogasification of solid and/or liquid carbonaceousmaterials to obtain a process gas which is steam pyrolized withadditional natural gas to form synthesis gas. The syn gas is convertedto methanol which may be carbonylated to acetic acid. The methodlikewise produces hydrogen which may be used in connection with thisinvention as noted above. U.S. Pat. No. 5,821,111, which discloses aprocess for converting waste biomass through gasification into synthesisgas as well as U.S. Pat. No. 6,685,754, the disclosures of which areincorporated herein by reference.

In one optional embodiment, the acetic acid fed to the hydrogenationreaction may also comprise other carboxylic acids and anhydrides, aswell as acetaldehyde and acetone. Preferably, a suitable acetic acidfeed stream comprises one or more of the compounds selected from thegroup consisting of acetic acid, acetic anhydride, acetaldehyde, andmixtures thereof. These other compounds may also be hydrogenated in theprocesses of the present invention. In some embodiments, the present ofcarboxylic acids, such as propanoic acid or its anhydride, may bebeneficial in producing propanol.

Alternatively, acetic acid in vapor form may be taken directly as crudeproduct from the flash vessel of a methanol carbonylation unit of theclass described in U.S. Pat. No. 6,657,078, the entirety of which isincorporated herein by reference. The crude vapor product, for example,may be fed directly to the ethanol synthesis reaction zones of thepresent invention without the need for condensing the acetic acid andlight ends or removing water, saving overall processing costs.

The acetic acid may be vaporized at the reaction temperature, followingwhich the vaporized acetic acid can be fed along with hydrogen in anundiluted state or diluted with a relatively inert carrier gas, such asnitrogen, argon, helium, carbon dioxide and the like. For reactions runin the vapor phase, the temperature should be controlled in the systemsuch that it does not fall below the dew point of acetic acid. In oneembodiment the acetic acid may be vaporized at the boiling point ofacetic acid at the particular pressure, and then the vaporized aceticacid may be further heated to the reactor inlet temperature. In anotherembodiment, the acetic acid is transferred to the vapor state by passinghydrogen, recycle gas, another suitable gas, or mixtures thereof throughthe acetic acid at a temperature below the boiling point of acetic acid,thereby humidifying the carrier gas with acetic acid vapors, followed byheating the mixed vapors up to the reactor inlet temperature.Preferably, the acetic acid is transferred to the vapor by passinghydrogen and/or recycle gas through the acetic acid at a temperature ator below 125° C., followed by heating of the combined gaseous stream tothe reactor inlet temperature.

In particular, the hydrogenation of acetic acid may achieve favorableconversion of acetic acid and favorable selectivity and productivity toethanol. For purposes of the present invention, the term “conversion”refers to the amount of acetic acid in the feed that is converted to acompound other than acetic acid. Conversion is expressed as a molepercentage based on acetic acid in the feed. The conversion may be atleast 10%, e.g., at least 20%, at least 40%, at least 50%, at least 60%,at least 70% or at least 80%. Although catalysts that have highconversions are desirable, such as at least 80% or at least 90%, in someembodiments a low conversion may be acceptable at high selectivity forethanol. It is, of course, well understood that in many cases, it ispossible to compensate for conversion by appropriate recycle streams oruse of larger reactors, but it is more difficult to compensate for poorselectivity.

Selectivity is expressed as a mole percent based on converted aceticacid. It should be understood that each compound converted from aceticacid has an independent selectivity and that selectivity is independentfrom conversion. For example, if 50 mole % of the converted acetic acidis converted to ethanol, we refer to the ethanol selectivity as 50%.Preferably, the catalyst selectivity to ethoxylates is at least 60%,e.g., at least 70%, or at least 80%. As used herein, the term“ethoxylates” refers specifically to the compounds ethanol,acetaldehyde, and ethyl acetate. Preferably, the selectivity to ethanolis at least 80%, e.g., at least 85% or at least 88%. Preferredembodiments of the hydrogenation process also have low selectivity toundesirable products, such as methane, ethane, and carbon dioxide. Theselectivity to these undesirable products preferably is less than 4%,e.g., less than 2% or less than 1%. More preferably, these undesirableproducts are not detectable. Formation of alkanes may be low, andideally less than 2%, less than 1%, or less than 0.5% of the acetic acidpassed over the catalyst is converted to alkanes, which have littlevalue other than as fuel.

The term “productivity,” as used herein, refers to the grams of aspecified product, e.g., ethanol, formed during the hydrogenation basedon the kilograms of catalyst used per hour. A productivity of at least200 grams of ethanol per kilogram catalyst per hour, e.g., at least 400grams of ethanol per kilogram catalyst per hour or at least 600 grams ofethanol per kilogram catalyst per hour, is preferred. In terms ofranges, the productivity preferably is from 200 to 3,000 grams ofethanol per kilogram catalyst per hour, e.g., from 400 to 2,500 perkilogram catalyst per hour or from 600 to 2,000 per kilogram catalystper hour.

In various embodiments, the crude ethanol product produced by thehydrogenation process, before any subsequent processing, such aspurification and separation, will typically comprise unreacted aceticacid, ethanol and water. As used herein, the term “crude ethanolproduct” refers to any composition comprising from 5 to 70 wt. % ethanoland from 5 to 35 wt. % water. In some exemplary embodiments, the crudeethanol product comprises ethanol in an amount from 5 wt. % to 70 wt. %,e.g., from 10 wt. % to 60 wt. %, or from 15 wt. % to 50 wt. %, based onthe total weight of the crude ethanol product. Preferably, the crudeethanol product contains at least 10 wt. % ethanol, at least 15 wt. %ethanol or at least 20 wt. % ethanol. The crude ethanol producttypically will further comprise unreacted acetic acid, depending onconversion, for example, in an amount of less than 90 wt. %, e.g., lessthan 80 wt. % or less than 70 wt. %. In terms of ranges, the unreactedacetic acid is preferably from 0 to 90 wt. %, e.g., from 5 to 80 wt. %,from 15 to 70 wt. %, from 20 to 70 wt. % or from 25 to 65 wt. %. Aswater is formed in the reaction process, water will generally be presentin the crude ethanol product, for example, in amounts ranging from 5 to35 wt. %, e.g., from 10 to 30 wt. % or from 10 to 26 wt. %. Ethylacetate may also be produced during the hydrogenation of acetic acid orthrough side reactions and may be present, for example, in amountsranging from 0 to 20 wt. %, e.g., from 0 to 15 wt. %, from 1 to 12 wt. %or from 3 to 10 wt. %. Acetaldehyde may also be produced through sidereactions and may be present, for example, in amounts ranging from 0 to10 wt. %, e.g., from 0 to 3 wt. %, from 0.1 to 3 wt. % or from 0.2 to 2wt. %. Other components, such as, for example, esters, ethers,aldehydes, ketones, alkanes, and carbon dioxide, if detectable,collectively may be present in amounts less than 10 wt. %, e.g., lessthan 6 wt. % or less than 4 wt. %. In terms of ranges, other componentsmay be present in an amount from 0.1 to 10 wt. %, e.g., from 0.1 to 6wt. %, or from 0.1 to 4 wt. %. Exemplary embodiments of crude ethanolcompositional ranges are provided in Table 1.

TABLE 1 CRUDE ETHANOL PRODUCT COMPOSITIONS Conc. Conc. Conc. Conc.Component (wt. %) (wt. %) (wt. %) (wt. %) Ethanol 5 to 70 10 to 60  15to 50 25 to 50 Acetic Acid 0 to 90 5 to 80 15 to 70 20 to 70 Water 5 to35 5 to 30 10 to 30 10 to 26 Ethyl Acetate 0 to 20 0 to 15  1 to 12  3to 10 Acetaldehyde 0 to 10 0 to 3  0.1 to 3   0.2 to 2   Others 0.1 to10   0.1 to 6   0.1 to 4   —Purification

FIGS. 1A and 1B show a hydrogenation system 100 suitable for thehydrogenation of acetic acid and separating ethanol from the crudereaction mixture according to one embodiment of the invention. System100 comprises reaction zone 101 and distillation zone 102. Reaction zone101 comprises reactor 103, hydrogen feed line 104 and acetic acid feedline 105. Distillation zone 102 comprises flasher 106, first column 107,second column 108, and third column 109. Hydrogen and acetic acid arefed to a vaporizer 110 via lines 104 and 105, respectively, to create avapor feed stream in line 111 that is directed to reactor 103. In oneembodiment, lines 104 and 105 may be combined and jointly fed to thevaporizer 110, e.g., in one stream containing both hydrogen and aceticacid. The temperature of the vapor feed stream in line 111 is preferablyfrom 100° C. to 350° C., e.g., from 120° C. to 310° C. or from 150° C.to 300° C. Any feed that is not vaporized is removed from vaporizer 110,as shown in FIG. 1A, and may be recycled thereto. In addition, althoughFIG. 1A shows line 111 being directed to the top of reactor 103, line111 may be directed to the side, upper portion, or bottom of reactor103. Further modifications and additional components to reaction zone101 are described below in FIG. 2.

Reactor 103 contains the catalyst that is used in the hydrogenation ofthe carboxylic acid, preferably acetic acid. In one embodiment, one ormore guard beds (not shown) may be used to protect the catalyst frompoisons or undesirable impurities contained in the feed orreturn/recycle streams. Such guard beds may be employed in the vapor orliquid streams. Suitable guard bed materials are known in the art andinclude, for example, carbon, silica, alumina, ceramic, or resins. Inone aspect, the guard bed media is functionalized to trap particularspecies such as sulfur or halogens. During the hydrogenation process, acrude ethanol product stream is withdrawn, preferably continuously, fromreactor 103 via line 112. The crude ethanol product stream may becondensed and fed to flasher 106, which, in turn, provides a vaporstream and a liquid stream. The flasher 106 in one embodiment preferablyoperates at a temperature of from 50° C. to 500° C., e.g., from 70° C.to 400° C. or from 100° C. to 350° C. In one embodiment, the pressure offlasher 106 preferably is from 50 kPa to 2000 kPa, e.g., from 75 kPa to1500 kPa or from 100 to 1000 kPa. In one preferred embodiment thetemperature and pressure of the flasher is similar to the temperatureand pressure of the reactor 103.

The vapor stream exiting the flasher 106 may comprise hydrogen andhydrocarbons, which may be purged and/or returned to reaction zone 101via line 113. As shown in FIG. 1A, the returned portion of the vaporstream passes through compressor 114 and is combined with the hydrogenfeed and co-fed to vaporizer 110.

The liquid from flasher 106 is withdrawn and pumped as a feedcomposition via line 115 to the side of first column 107, also referredto as the acid separation column. The contents of line 115 typicallywill be substantially similar to the product obtained directly from thereactor, and may, in fact, also be characterized as a crude ethanolproduct. However, the feed composition in line 115 preferably hassubstantially no hydrogen, carbon dioxide, methane or ethane, which areremoved by flasher 106. Exemplary components of liquid in line 115 areprovided in Table 2. It should be understood that liquid line 115 maycontain other components, not listed, such as components in the feed.

TABLE 2 FEED COMPOSITION Conc. (wt. %) Conc. (wt. %) Conc. (wt. %)Ethanol 5 to 70 10 to 60  15 to 50 Acetic Acid <90 5 to 80 15 to 70Water 5 to 35 5 to 30 10 to 30 Ethyl Acetate <20 0.001 to 15     1 to 12Acetaldehyde <10 0.001 to 3    0.1 to 3   Acetal <5 0.001 to 2    0.005to 1    Acetone <5 0.0005 to 0.05   0.001 to 0.03  Other Esters <5<0.005 <0.001 Other Ethers <5 <0.005 <0.001 Other Alcohols <5 <0.005<0.001

The amounts indicated as less than (<) in the tables throughout presentapplication are preferably not present and if present may be present intrace amounts or in amounts greater than 0.0001 wt. %.

The “other esters” in Table 2 may include, but are not limited to, ethylpropionate, methyl acetate, isopropyl acetate, n-propyl acetate, n-butylacetate or mixtures thereof. The “other ethers” in Table 2 may include,but are not limited to, diethyl ether, methyl ethyl ether, isobutylethyl ether or mixtures thereof. The “other alcohols” in Table 2 mayinclude, but are not limited to, methanol, isopropanol, n-propanol,n-butanol or mixtures thereof. In one embodiment, the feed composition,e.g., line 115, may comprise propanol, e.g., isopropanol and/orn-propanol, in an amount from 0.001 to 0.1 wt. %, from 0.001 to 0.05 wt.% or from 0.001 to 0.03 wt. %. In should be understood that these othercomponents may be carried through in any of the distillate or residuestreams described herein and will not be further described herein,unless indicated otherwise.

When the content of acetic acid in line 115 is less than 5 wt. %, theacid separation column 107 may be skipped and line 115 may be introduceddirectly to second column 108, also referred to herein as a light endscolumn.

In the embodiment shown in FIG. 1A, line 115 is introduced in the lowerpart of first column 107, e.g., lower half or lower third. In firstcolumn 107, unreacted acetic acid, a portion of the water, and otherheavy components, if present, are removed from the composition in line115 and are withdrawn, preferably continuously, as residue. Some or allof the residue may be returned and/or recycled back to reaction zone 101via line 116. First column 107 also forms an overhead distillate, whichis withdrawn in line 117, and which may be condensed and refluxed, forexample, at a ratio of from 10:1 to 1:10, e.g., from 3:1 to 1:3 or from1:2 to 2:1.

Any of columns 107, 108 or 109 may comprise any distillation columncapable of separation and/or purification. The columns preferablycomprise tray columns having from 1 to 150 trays, e.g., from 10 to 100trays, from 20 to 95 trays or from 30 to 75 trays. The trays may besieve trays, fixed valve trays, movable valve trays, or any othersuitable design known in the art. In other embodiments, a packed columnmay be used. For packed columns, structured packing or random packingmay be employed. The trays or packing may be arranged in one continuouscolumn or they may be arranged in two or more columns such that thevapor from the first section enters the second section while the liquidfrom the second section enters the first section, etc.

The associated condensers and liquid separation vessels that may beemployed with each of the distillation columns may be of anyconventional design and are simplified in FIGS. 1A and 1B. As shown inFIGS. 1A and 1B, heat may be supplied to the base of each column or to acirculating bottom stream through a heat exchanger or reboiler. Othertypes of reboilers, such as internal reboilers, may also be used in someembodiments. The heat that is provided to reboilers may be derived fromany heat generated during the process that is integrated with thereboilers or from an external source such as another heat generatingchemical process or a boiler. Although one reactor and one flasher areshown in FIGS. 1A and 1B, additional reactors, flashers, condensers,heating elements, and other components may be used in embodiments of thepresent invention. As will be recognized by those skilled in the art,various condensers, pumps, compressors, reboilers, drums, valves,connectors, separation vessels, etc., normally employed in carrying outchemical processes may also be combined and employed in the processes ofthe present invention.

The temperatures and pressures employed in any of the columns may vary.As a practical matter, pressures from 10 kPa to 3000 kPa will generallybe employed in these zones although in some embodiments subatmosphericpressures may be employed as well as superatmospheric pressures.Temperatures within the various zones will normally range between theboiling points of the composition removed as the distillate and thecomposition removed as the residue. It will be recognized by thoseskilled in the art that the temperature at a given location in anoperating distillation column is dependent on the composition of thematerial at that location and the pressure of column. In addition, feedrates may vary depending on the size of the production process and, ifdescribed, may be generically referred to in terms of feed weightratios.

When column 107 is operated under standard atmospheric pressure, thetemperature of the residue exiting in line 116 from column 107preferably is from 95° C. to 120° C., e.g., from 105° C. to 117° C. orfrom 110° C. to 115° C. The temperature of the distillate exiting inline 117 from column 107 preferably is from 70° C. to 110° C., e.g.,from 75° C. to 95° C. or from 80° C. to 90° C. In other embodiments, thepressure of first column 107 may range from 0.1 kPa to 510 kPa, e.g.,from 1 kPa to 475 kPa or from 1 kPa to 375 kPa. Exemplary components ofthe distillate and residue compositions for first column 107 areprovided in Table 3 below. It should also be understood that thedistillate and residue may also contain other components, not listed,such as components in the feed. For convenience, the distillate andresidue of the first column may also be referred to as the “firstdistillate” or “first residue.” The distillates or residues of the othercolumns may also be referred to with similar numeric modifiers (second,third, etc.) in order to distinguish them from one another, but suchmodifiers should not be construed as requiring any particular separationorder.

TABLE 3 FIRST COLUMN Conc. (wt. %) Conc. (wt. %) Conc. (wt. %)Distillate Ethanol 20 to 75 30 to 70 40 to 65 Water 10 to 40 15 to 35 20to 35 Acetic Acid <2 0.001 to 0.5  0.01 to 0.2  Ethyl Acetate <60 5.0 to40  10 to 30 Acetaldehyde <10 0.001 to 5 0.01 to 4   Acetal <0.1 <0.1<0.05 Acetone <0.05 0.001 to 0.03   0.01 to 0.025 Residue Acetic Acid 60 to 100 70 to 95 85 to 92 Water <30  1 to 20  1 to 15 Ethanol <1 <0.9<0.07

As shown in Table 3, without being bound by theory, it has surprisinglyand unexpectedly been discovered that when any amount of acetal isdetected in the feed that is introduced to the acid separation column(first column 107), the acetal appears to decompose in the column suchthat less or even no detectable amounts are present in the distillateand/or residue.

Depending on the reaction conditions, the crude ethanol product exitingreactor 103 in line 112 may comprise ethanol, acetic acid (unconverted),ethyl acetate, and water. After exiting reactor 103, a non-catalyzedequilibrium reaction may occur between the components contained in thecrude ethanol product until it is added to flasher 106 and/or firstcolumn 107. This equilibrium reaction tends to drive the crude ethanolproduct to an equilibrium between ethanol/acetic acid and ethylacetate/water, as shown below.EtOH+HOAc

EtOAc+H₂O

In the event the crude ethanol product is temporarily stored, e.g., in aholding tank, prior to being directed to distillation zone 102, extendedresidence times may be encountered. Generally, the longer the residencetime between reaction zone 101 and distillation zone 102, the greaterthe formation of ethyl acetate. For example, when the residence timebetween reaction zone 101 and distillation zone 102 is greater than 5days, significantly more ethyl acetate may form at the expense ofethanol. Thus, shorter residence times between reaction zone 101 anddistillation zone 102 are generally preferred in order to maximize theamount of ethanol formed. In one embodiment, a holding tank (not shown),is included between the reaction zone 101 and distillation zone 102 fortemporarily storing the liquid component from line 115 for up to 5 days,e.g., up to 1 day, or up to 1 hour. In a preferred embodiment no tank isincluded and the condensed liquids are fed directly to the firstdistillation column 107. In addition, the rate at which thenon-catalyzed reaction occurs may increase as the temperature of thecrude ethanol product, e.g., in line 115, increases. These reactionrates may be particularly problematic at temperatures exceeding 30° C.,e.g., exceeding 40° C. or exceeding 50° C. Thus, in one embodiment, thetemperature of liquid components in line 115 or in the optional holdingtank is maintained at a temperature less than 40° C., e.g., less than30° C. or less than 20° C. One or more cooling devices may be used toreduce the temperature of the liquid in line 115.

As discussed above, a holding tank (not shown) may be included betweenthe reaction zone 101 and distillation zone 102 for temporarily storingthe liquid component from line 115, for example from 1 to 24 hours,optionally at a temperature of about 21° C., and corresponding to anethyl acetate formation of from 0.01 wt. % to 1.0 wt. % respectively. Inaddition, the rate at which the non-catalyzed reaction occurs mayincrease as the temperature of the crude ethanol product is increased.For example, as the temperature of the crude ethanol product in line 115increase from 4° C. to 21° C., the rate of ethyl acetate formation mayincrease from about 0.01 wt. % per hour to about 0.005 wt. % per hour.Thus, in one embodiment, the temperature of liquid components in line115 or in the optional holding tank is maintained at a temperature lessthan 21° C., e.g., less than 4° C. or less than −10° C.

In addition, it has now been discovered that the above-describedequilibrium reaction may also favor ethanol formation in the top regionof first column 107.

The distillate, e.g., overhead stream, of column 107 optionally iscondensed and refluxed as shown in FIG. 1A, preferably, at a refluxratio of 1:5 to 10:1. The distillate in line 117 preferably comprisesethanol, ethyl acetate, and water, along with other impurities, whichmay be difficult to separate due to the formation of binary and tertiaryazeotropes.

The first distillate in line 117 is introduced to the second column 108,also referred to as the “light ends column,” preferably in the middlepart of column 108, e.g., middle half or middle third. As one example,when a 25 tray column is utilized in a column without water extraction,line 117 is introduced at tray 17. In one embodiment, the second column108 may be an extractive distillation column. In such embodiments, anextraction agent, such as water, may be added to second column 108. Ifthe extraction agent comprises water, it may be obtained from anexternal source or from an internal return/recycle line from one or moreof the other columns.

Second column 108 may be a tray column or packed column. In oneembodiment, second column 108 is a tray column having from 5 to 70trays, e.g., from 15 to 50 trays or from 20 to 45 trays.

Although the temperature and pressure of second column 108 may vary,when at atmospheric pressure the temperature of the second residueexiting in line 118 from second column 108 preferably is from 60° C. to90° C., e.g., from 70° C. to 90° C. or from 80° C. to 90° C. Thetemperature of the second distillate exiting in line 120 from secondcolumn 108 preferably is from 50° C. to 90° C., e.g., from 60° C. to 80°C. or from 60° C. to 70° C. Column 108 may operate at atmosphericpressure. In other embodiments, the pressure of second column 108 mayrange from 0.1 kPa to 510 kPa, e.g., from 1 kPa to 475 kPa or from 1 kPato 375 kPa. Exemplary components for the distillate and residuecompositions for second column 108 are provided in Table 4 below. Itshould be understood that the distillate and residue may also containother components, not listed, such as components in the feed.

TABLE 4 SECOND COLUMN Conc. (wt. %) Conc. (wt. %) Conc. (wt. %)Distillate Ethyl Acetate 10 to 90 25 to 90 50 to 90 Acetaldehyde  1 to25  1 to 15 1 to 8 Water  1 to 25  1 to 20  4 to 16 Ethanol <30 0.001 to15   0.01 to 5   Acetal <5 0.001 to 2    0.01 to 1   Residue Water 30 to70 30 to 60 30 to 50 Ethanol 20 to 75 30 to 70 40 to 70 Ethyl Acetate <30.001 to 2    0.001 to 0.5  Acetic Acid <0.5 0.001 to 0.3  0.001 to 0.2 

The weight ratio of ethanol in the second residue to ethanol in thesecond distillate preferably is at least 3:1, e.g., at least 6:1, atleast 8:1, at least 10:1 or at least 15:1. The weight ratio of ethylacetate in the second residue to ethyl acetate in the second distillatepreferably is less than 0.4:1, e.g., less than 0.2:1 or less than 0.1:1.In embodiments that use an extractive column with water as an extractionagent as the second column 108, the weight ratio of ethyl acetate in thesecond residue to ethyl acetate in the second distillate approacheszero.

As shown, the second residue from the bottom of second column 108, whichcomprises ethanol and water, is fed via line 118 to third column 109,also referred to as the “product column.” More preferably, the secondresidue in line 118 is introduced in the lower part of third column 109,e.g., lower half or lower third. Third column 109 recovers ethanol,which preferably is substantially pure other than the azeotropic watercontent, as the distillate in line 119. The distillate of third column109 preferably is refluxed as shown in FIG. 1A, for example, at a refluxratio of from 1:10 to 10:1, e.g., from 1:3 to 3:1 or from 1:2 to 2:1.The third residue in line 121, which preferably comprises primarilywater, preferably is removed from the system 100 or may be partiallyreturned to any portion of the system 100. Third column 109 ispreferably a tray column as described above and preferably operates atatmospheric pressure. The temperature of the third distillate exiting inline 119 from third column 109 preferably is from 60° C. to 110° C.,e.g., from 70° C. to 100° C. or from 75° C. to 95° C. The temperature ofthe third residue exiting from third column 109 preferably is from 70°C. to 115° C., e.g., from 80° C. to 110° C. or from 85° C. to 105° C.,when the column is operated at atmospheric pressure. Exemplarycomponents of the distillate and residue compositions for third column109 are provided in Table 5 below. It should be understood that thedistillate and residue may also contain other components, not listed,such as components in the feed.

TABLE 5 THIRD COLUMN Conc. (wt. %) Conc. (wt. %) Conc. (wt. %)Distillate Ethanol 75 to 96 80 to 96 85 to 96 Water <12 1 to 9 3 to 8Acetic Acid <1 0.001 to 0.1  0.005 to 0.01  Ethyl Acetate <5 0.001 to4    0.01 to 3   Residue Water  75 to 100  80 to 100  90 to 100 Ethanol<0.8 0.001 to 0.5  0.005 to 0.05  Ethyl Acetate <1 0.001 to 0.5  0.005to 0.2  Acetic Acid <2 0.001 to 0.5  0.005 to 0.2 

Any of the compounds that are carried through the distillation processfrom the feed or crude reaction product generally remain in the thirddistillate in amounts of less 0.1 wt. %, based on the total weight ofthe third distillate composition, e.g., less than 0.05 wt. % or lessthan 0.02 wt. %. In one embodiment, one or more side streams may removeimpurities from any of the columns 107, 108 and/or 109 in the system100. Preferably at least one side stream is used to remove impuritiesfrom the third column 109. The impurities may be purged and/or retainedwithin the system 100.

The third distillate in line 119 may be further purified to form ananhydrous ethanol product stream, i.e., “finished anhydrous ethanol,”using one or more additional separation systems, such as, for example,distillation columns (e.g., a finishing column) or molecular sieves.

Returning to second column 108, the distillate in line 120 preferably isrefluxed as shown in FIG. 1A, for example, at a reflux ratio of from1:10 to 10:1, e.g., from 1:5 to 5:1 or from 1:3 to 3:1. The distillatefrom second column 108 may be purged. Alternatively, since it containsethyl acetate, all or a portion of the distillate from second column 108may be recycled to reaction zone 101 via line 120 in order to convertthe ethyl acetate to additional ethanol. As shown in FIG. 1B, all or aportion the distillate may be recycled to reactor 103, as shown by line120, and may be co-fed with the acetic acid feed line 105. In anotherembodiments, the second distillate in line 120 may be further purifiedto remove impurities, such as acetaldehyde, using one or more additionalcolumns (not shown).

Although one reactor and one flasher are shown in FIGS. 1A and 1B,additional reactors and/or components may be included in variousoptional embodiments of the present invention. FIG. 2 represents ahydrogenation system 100′ that comprises dual reactors 103, 103′, dualflashers 106, 106′, heat exchanger 130, and pre-heater 131. In thisembodiment, acetic acid in line 105, along with the recycled acetic acidin line 116 and optionally recycled components from line 120, are heatedin a heat exchanger 130 and sent to vaporizer 110 via line 132. Thetemperature of the contents of line 132 preferably is from 30° C. to150° C., e.g., from 50° C. to 130° C. or from 75° C. to 120° C. Hydrogenis fed via line 104 to vaporizer 110, which forms vaporized stream 111.Vaporized stream 111 passes through pre-heater 131, which further heatsstream 111 to a temperature of preferably from 200° C. to 300° C., e.g.,from 210° C. to 275° C. or from 220° C. to 260° C. The heated stream isthen fed to first reactor 103. In order to control the reactionexotherm, the crude reaction mixture is removed from first reactor 103via line 133 and cooled before being fed to a second reactor 103′, suchthat the temperature of the reactants and products in contact with thecatalyst is maintained at or below 310° C. in order to minimize theformation of undesired byproducts including methane, ethane, carbondioxide, and/or carbon monoxide. Additionally, above about 320° C.corrosion can become severe necessitating the use of exotic andexpensive alloy materials. The temperature of the contents in line 133after cooling preferably is from 200° C. to 300° C., e.g., from 210° C.to 275° C. or from 220° C. to 260° C. The reactors 103 and 103′ may bethe same size and configuration or they may be of different size andconfiguration. Each reactor preferably contains the same type ofcatalyst, although additional and/or different catalysts may be used foreach reactor. As an example, the catalysts mentioned above may beutilized. Also, mixtures of catalysts, mixtures of catalysts and inertmaterials, and/or catalysts with differing active metal concentrationsmay be utilized. For example, the catalyst may include the same types ofmetal in varying metal ratios. A crude ethanol product stream iswithdrawn, preferably continuously, from reactor 103′ via line 112 andpasses as a heating medium through heat exchanger 130 before beingcondensed and fed to first flasher 106. Thus, heat from the crudeethanol product stream advantageously may be employed to preheat theacetic acid feed prior to its introduction into vaporizer 110.Conversely, the acetic acid feed may be used as a cooling medium to coolthe crude ethanol product stream prior to its introduction to firstflasher 106. The vapor stream exiting the first flasher compriseshydrogen and hydrocarbons, which may be purged and/or returned toreaction zone 101 via line 113. As shown in FIG. 2, at least a portionof the recycled vapor stream passes through compressor 114 and is co-fedwith the hydrogen (or combined with hydrogen and then co-fed) tovaporizer 110.

The remaining liquid in flasher 106 is withdrawn via line 134 and fed toa second flasher 106′ to remove any residual vapor that is dissolved inthe liquid. Second flasher 106′ may operate at a lower temperatureand/or pressure than the first flasher 106. In one embodiment, thetemperature of second flasher 106′ preferably is from 20° C. to 100° C.,e.g., from 30° C. to 85° C. or from 40° C. to 70° C. In one embodiment,the temperature of second flasher 106′ preferably is at least 50° C.lower than first flasher 106, e.g., at least 75° C. lower or at least100° C. lower. The pressure of second flasher 106′ preferably is from0.1 kPa to 1000 kPa, e.g., from 0.1 kPa to 500 kPa or from 0.1 kPa to100 kPa. In one embodiment, the pressure of second flasher 106′preferably is at least 50 kPa lower than first flasher 106, e.g., atleast 100 kPa lower or at least 200 kPa lower. The vapor stream 135exiting the second flasher may comprise hydrogen and hydrocarbons, whichmay be purged and/or returned to the reaction zone in a manner similarto that of the first flasher 106. The remaining liquid in flasher 106′is withdrawn and pumped via line 115 to the side of the first column(not shown in FIG. 2) and is further purified to form an ethanol productstream, i.e., “finished ethanol,” as described, for example, inconnection with FIGS. 1A and 1B.

Finished Ethanol

The finished ethanol composition obtained by the processes of thepresent invention preferably comprises from 75 to 96 wt. % ethanol,e.g., from 80 to 96 wt. % or from 85 to 96 wt. % ethanol, based on thetotal weight of the finished ethanol composition. Exemplary finishedethanol compositional ranges are provided below in Table 6.

TABLE 6 FINISHED ETHANOL COMPOSITIONS Component Conc. (wt. %) Conc. (wt.%) Conc. (wt. %) Ethanol 75 to 96 80 to 96 85 to 96 Water <12 1 to 9 3to 8 Acetic Acid <1 <0.1 <0.01 Ethyl Acetate <2 <0.5 <0.05 Acetal <0.05<0.01 <0.005 Acetone <0.05 <0.01 <0.005 Isopropanol <0.5 <0.1 <0.05n-propanol <0.5 <0.1 <0.05

The finished ethanol composition of the present invention preferablycontains very low amounts, e.g., less than 0.5 wt. %, of other alcohols,such as methanol, butanol, isobutanol, isoamyl alcohol and other C₄-C₂₀alcohols. In one embodiment, the amount of isopropanol in the finishedethanol is from 95 to 1,000 wppm, e.g., from 100 to 700 wppm, or from150 to 500 wppm. In one embodiment, the finished ethanol compositionpreferably is substantially free of acetaldehyde and may comprise lessthan 8 wppm of acetaldehyde, e.g., less than 5 wppm or less than 1 wppm.

The finished ethanol composition produced by the embodiments of thepresent invention may be used in a variety of applications includingfuels, solvents, chemical feedstocks, pharmaceutical products,cleansers, sanitizers, hydrogenation transport or consumption. In fuelapplications, the finished ethanol composition may be blended withgasoline for motor vehicles such as automobiles, boats and small pistonengine aircrafts. In non-fuel applications, the finished ethanolcomposition may be used as a solvent for toiletry and cosmeticpreparations, detergents, disinfectants, coatings, inks, andpharmaceuticals. The finished ethanol composition may also be used as aprocessing solvent in manufacturing processes for medicinal products,food preparations, dyes, photochemicals and latex processing.

The finished ethanol composition may also be used a chemical feedstockto make other chemicals such as vinegar, ethyl acrylate, ethyl acetate,ethylene, glycol ethers, ethylamines, aldehydes, and higher alcohols,especially butanol. In the production of ethyl acetate, the finishedethanol composition may be esterified with acetic acid or reacted withpolyvinyl acetate. The finished ethanol composition may be dehydrated toproduce ethylene. Any of known dehydration catalysts can be employed into dehydrate ethanol, such as those described in copending U.S.application Ser. No. 12/221,137 and U.S. application Ser. No.12/221,138, the entire contents and disclosures of which are herebyincorporated by reference. A zeolite catalyst, for example, may beemployed as the dehydration catalyst. Preferably, the zeolite has a porediameter of at least about 0.6 nm, and preferred zeolites includedehydration catalysts selected from the group consisting of mordenites,ZSM-5, a zeolite X and a zeolite Y. Zeolite X is described, for example,in U.S. Pat. No. 2,882,244 and zeolite Yin U.S. Pat. No. 3,130,007, theentireties of which are hereby incorporated by reference.

In order that the invention disclosed herein may be more efficientlyunderstood, an example is provided below. The following examplesdescribe the various distillation processes of the present invention.

EXAMPLES Example 1

A crude ethanol product comprising ethanol, acetic acid, water and ethylacetate was produced by reacting a vaporized feed comprising 95.2 wt. %acetic acid and 4.6 wt. % water with hydrogen in the presence of acatalyst comprising 1.6 wt. % platinum and 1 wt. % tin supported on ⅛inch calcium silicate modified silica extrudates at an averagetemperature of 291° C., an outlet pressure of 2,063 kPa. Unreactedhydrogen was recycled back to the inlet of the reactor such that thetotal H₂/acetic acid molar ratio was 5.8 at a GHSV of 3,893 hr⁻¹. Underthese conditions, 42.8% of the acetic acid was converted, and theselectivity to ethanol was 87.1%, selectivity to ethyl acetate was 8.4%,and selectivity to acetaldehyde was 3.5%. The crude ethanol product waspurified using a separation scheme having distillation columns as shownin FIG. 1A.

The crude ethanol product was fed to the first column at a feed rate of20 g/min. The composition of the liquid feed is provided in Table 7. Thefirst column is a 2 inch diameter Oldershaw with 50 trays. The columnwas operated at a temperature of 115° C. at atmospheric pressure. Unlessotherwise indicated, a column operating temperature is the temperatureof the liquid in the reboiler and the pressure at the top of the columnis atmospheric (approximately one atmosphere). The column differentialpressure between the trays in the first column was 7.4 kPa. The firstresidue was withdrawn at a flow rate of 12.4 g/min and returned to thehydrogenation reactor.

The first distillate was condensed and refluxed at a 1:1 ratio at thetop of the first column, and a portion of the distillate was introducedto the second column at a feed rate of 7.6 g/min. The second column is a2 inch diameter Oldershaw design equipped with 25 trays. The secondcolumn was operated at a temperature of 82° C. at atmospheric pressure.The column differential pressure between the trays in the second columnwas 2.6 kPa. The second residue was withdrawn at a flow rate of 5.8g/min and directed to the third column. The second distillate wasrefluxed at a ratio of 4.5:0.5 and the remaining distillate wascollected for analysis. The compositions of the feed, distillates, andresidues are provided in Table 7.

TABLE 7 First Column Second Column Feed Distillate Residue DistillateResidue Component (wt. %) (wt. %) (wt. %) (wt. %) (wt. %) Water 13.824.7 5.6 5.1 30.8 Acetaldehyde nd 1.8 nd 8.3 nd Acetic Acid 55.0 0.0893.8 0.03 0.1 Ethanol 23.4 57.6 0.06 12.4 67.6 Ethyl Acetate 6.5 15.1 nd76.0 nd Acetal 0.7 0.1 nd 0.006 0.03 Acetone nd 0.01 nd 0.03 nd

Residue from the second column was collected from several runs andintroduced above the 25 tray to the third column, a 2 inch Oldershawcontaining 60 trays, at a rate of 10 g/min. The third column wasoperated at a temperature of 103° C. at standard pressure. The columndifferential pressure between the trays in the third column was 6.2 kPa.The third residue was withdrawn at a flow rate of 2.7 g/min. The thirddistillate was condensed and refluxed at a 3:1 ratio at the top of thethird column, and recovered an ethanol composition as shown in Table 8.The ethanol composition also contained 10 ppm of n-butyl acetate.

TABLE 8 Third Column Feed Distillate Residue Component (wt. %) (wt. %)(wt. %) Acetic Acid 0.098 0.001 0.4 Ethanol 65.7 93.8 0.004 Water 35.56.84 98 Ethyl Acetate 1.37 1.8 — Acetal 0.02 0.03 — Isopropanol 0.0040.005 — n-propanol 0.01 0 —

Example 2

A crude ethanol product comprising ethanol, acetic acid, water and ethylacetate was produced by reacting a vaporized feed comprising 96.3 wt. %acetic acid and 4.3 wt. % water with hydrogen in the presence of acatalyst comprising 1.6 wt. % platinum and 1% tin supported on ⅛ inchcalcium silicate modified silica extrudates at an average temperature of290° C., an outlet pressure of 2,049 kPa. Unreacted hydrogen wasrecycled back to the inlet of the reactor such that the total H₂/aceticacid molar ratio was 10.2 at a GHSV of 1,997 hr⁻¹. Under theseconditions, 74.5% of the acetic acid was converted, and the selectivityto ethanol was 87.9%, selectivity to ethyl acetate was 9.5%, andselectivity to acetaldehyde was 1.8%. The crude ethanol product waspurified using a separation scheme having distillation columns as shownin FIG. 1A.

The crude ethanol product was fed to the first column at a feed rate of20 g/min. The composition of the liquid feed is provided in Table 9. Thefirst column is a 2 inch diameter Oldershaw with 50 trays. The columnwas operated at a temperature of 116° C. at atmospheric pressure. Thecolumn differential pressure between the trays in the first column was8.1 kPa. The first residue was withdrawn at a flow rate of 10.7 g/minand returned to the hydrogenation reactor.

The first distillate was condensed and refluxed at a 1:1 ratio at thetop of the first column, and a portion of the distillate was introducedto the second column at a feed rate of 9.2 g/min. The second column is a2 inch diameter Oldershaw design equipped with 25 trays. The secondcolumn was operated at a temperature of 82° C. at atmospheric pressure.The column differential pressure between the trays in the second columnwas 2.4 kPa. The second residue was withdrawn at a flow rate of 7.1g/min and directed to the third column. The second distillate wasrefluxed at a ratio of 4.5:0.5 and the remaining distillate wascollected for analysis. The compositions of the feed, distillates, andresidues are provided in Table 9.

TABLE 9 First Column Second Column Feed Distillate Residue DistillateResidue Component (wt. %) (wt. %) (wt. %) (wt. %) (wt. %) Water 14.627.2 3.7 3.0 36.2 Acetaldehyde nd 1.5 nd 10.3 nd Acetic Acid 49.1 0.298.2 0.04 0.3 Ethanol 27.6 54.5 0.04 13.3 64.4 Ethyl Acetate 7.9 15.2 nd75.7 1.8 Acetal 0.7 0.1 nd 0.01 0.02 Acetone nd 0.01 nd 0.03 nd

Residue from the second column was collected from several runs andintroduced above the 25 tray to the third column, a 2 inch Oldershawcontaining 60 trays, at a rate of 10 g/min. The third column wasoperated at a temperature of 103° C. at standard pressure. The columndifferential pressure between the trays in the third column was 6.5 kPa.The third residue was withdrawn at a flow rate of 2.8 g/min. The thirddistillate was condensed and refluxed at a 3:1 ratio at the top of thethird column, and recovered an ethanol composition as shown in Table 10.The ethanol composition also contained 86 ppm of isopropanol and 2.3 ppmof n-propyl acetate.

TABLE 10 Third Column Feed Distillate Residue Component (wt. %) (wt. %)(wt. %) Acetic Acid 0.16 0.0028 0.77 Ethanol 64.4 92.3 0.8 Water 35.86.3 98.0 Ethyl Acetate 0.9 0.45 0.0007

While the invention has been described in detail, modifications withinthe spirit and scope of the invention will be readily apparent to thoseof skill in the art. In view of the foregoing discussion, relevantknowledge in the art and references discussed above in connection withthe Background and Detailed Description, the disclosures of which areall incorporated herein by reference. In addition, it should beunderstood that aspects of the invention and portions of variousembodiments and various features recited below and/or in the appendedclaims may be combined or interchanged either in whole or in part. Inthe foregoing descriptions of the various embodiments, those embodimentswhich refer to another embodiment may be appropriately combined withother embodiments as will be appreciated by one of skill in the art.Furthermore, those of ordinary skill in the art will appreciate that theforegoing description is by way of example only, and is not intended tolimit the invention.

1. A process for recovering ethanol, comprising: hydrogenating an aceticacid feed stream with excess hydrogen in a reactor in the presence of acatalyst to form a crude ethanol product; separating at least a portionof the crude ethanol product in a first flasher into a first vaporstream and an intermediate stream, wherein the first flasher is operatedat a pressure ranging from 50 kPa to 5000 kPa; separating at least aportion of the intermediate stream in a second flasher into a secondvapor stream and a liquid stream, wherein the second flasher is operatedat a pressure ranging from 0.1 kPa to 1000 kPa; and recovering ethanolfrom the liquid stream; wherein the second vapor stream compriseshydrogen and hydrocarbons and further wherein the liquid streamcomprises substantially no hydrogen, carbon dioxide, methane or ethane;and further wherein the first flasher is operated at higher pressurethan the second flasher.
 2. The process of claim 1, further comprisingseparating at least a portion of the liquid stream in a first columninto a first distillate comprising ethanol, water and ethyl acetate, anda first residue comprising acetic acid; separating at least a portion ofthe first distillate in a second column into a second distillatecomprising ethyl acetate and a second residue comprising ethanol andwater; and separating at least a portion of the second residue in athird column into a third distillate comprising ethanol and a thirdresidue comprising water.
 3. The process of claim 1, wherein the firstflasher is operated at a temperature ranging from 50° C. to 500° C. andwherein the second flasher is operated at a temperature ranging from 20°C. to 100° C.
 4. The process of claim 1, wherein the pressure of thesecond flasher is at least 50 kPa lower than the pressure of the firstflasher.
 5. The process of claim 1, wherein the temperature of thesecond flasher is at least 50° C. lower than the temperature of thefirst flasher.
 6. The process of claim 1, further comprising recycling aportion of the first vapor stream to the reactor.
 7. The process ofclaim 1, further comprising recycling a portion of the second vaporstream to the reactor.
 8. The process of claim 1, wherein the aceticacid is formed from methanol and carbon monoxide, wherein each of themethanol, the carbon monoxide, and hydrogen for the hydrogenating stepis derived from syngas, and wherein the syngas is derived from a carbonsource selected from the group consisting of natural gas, oil,petroleum, coal, biomass, and combinations thereof.
 9. A process forrecovering ethanol, comprising: providing a crude ethanol productcomprising ethanol, water, ethyl acetate, and acetaldehyde; separatingat least a portion of the crude ethanol product in a first flasher intoa first vapor stream and an intermediate stream, wherein the firstflasher is operated at a pressure of 50 kPa to 5000 kP; separating atleast a portion of the intermediate stream in a second flasher into asecond vapor stream and a liquid stream, wherein the second flasher isoperated at a pressure ranging from 0.1 kPa to 1000 kPa; and recoveringethanol from the liquid stream; wherein the second vapor streamcomprises hydrogen and hydrocarbons and further wherein the liquidstream comprises substantially no hydrogen carbon dioxide, methane orethane; and further wherein the first flasher is at higher pressure thanthe second flasher.
 10. The process of claim 9, wherein the crudeethanol product further comprises acetic and the process furthercomprises: separating at least a portion of the liquid stream in a firstcolumn into a first distillate comprising ethanol, water and ethylacetate, and a first residue comprising acetic acid; separating at leasta portion of the first distillate in a second column into a seconddistillate comprising ethyl acetate and a second residue comprisingethanol and water; and separating at least a portion of the secondresidue in a third column into a third distillate comprising ethanol anda third residue comprising water.
 11. The process of claim 9, whereinthe first flasher is operated at a temperature ranging from 50° C. to500° C. and wherein the second flasher is operated at a temperatureranging from 20° C. to 100° C.
 12. The process of claim 9, wherein thepressure of the second flasher is at least 50 kPa lower than thepressure of the first flasher.
 13. The process of claim 9, wherein thetemperature of the second flasher is at least 50° C. lower than thetemperature of the first flasher.
 14. The process of claim 9, furthercomprising recycling a portion of the first vapor stream to the reactor.15. The process of claim 9, further comprising recycling a portion ofthe second vapor stream to the reactor.
 16. The process of claim 9,wherein the acetic acid is formed from methanol and carbon monoxide,wherein each of the methanol, the carbon monoxide, and hydrogen for thehydrogenating step is derived from syngas, and wherein the syngas isderived from a carbon source selected from the group consisting ofnatural gas, oil, petroleum, coal, biomass, and combinations thereof.17. A process for recovering ethanol, comprising: hydrogenating anacetic acid feed stream with excess hydrogen in a reactor in thepresence of a catalyst to form a crude ethanol product; separating atleast a portion of the crude ethanol product in a first flasher into afirst vapor stream and an intermediate stream; separating at least aportion of the intermediate stream in a second flasher into a secondvapor stream and a liquid stream; and recovering ethanol from the liquidstream; wherein the second vapor stream comprises hydrogen andhydrocarbons and further wherein the liquid stream comprisessubstantially no hydrogen, carbon dioxide, methane or ethane; andfurther wherein the pressure of the second flasher is at least 50 kPalower than the pressure of the first flasher.